Process for preparing benzene

ABSTRACT

The present invention relates to a process for the endothermic, catalytic gas phase reaction of naphtha with hydrogen to form benzene, in which the reaction is carried out in 5 to 12 serial reaction zones under adiabatic conditions.

The present invention relates to a process for the endothermic,catalytic gas phase reaction of naphtha with hydrogen to form benzene,in which the reaction is carried out in 5 to 12 serial reaction zonesunder adiabatic conditions.

Naphtha is an untreated petroleum distillate from the refining ofpetroleum or natural gas, and one of the products typically recoveredfrom it is benzene. Benzene in turn is a key starting material for manyfurther petrochemicals.

For instance, benzene is used in the Chemical Industry for the synthesisof numerous compounds, such as, for example, aniline, styrene, nylon,synthetic rubber, plastics, detergents, insecticides, dyes and numerousfurther substances. Also obtained, by substitution, are numerousaromatics, such as, for example, phenol, nitrobenzene, aniline,chlorobenzene, hydroquinone and picric acid.

A further product, though one which has now taken a back seat forenvironmental reasons, is the use of the benzene as a fuel for internalcombustion engines operating in accordance with the Otto cycle.

The key reactions in the preparation of benzene from naphtha are set outin the formulae below (I to IV). Formula (I) relates to the conversionof cyclohexane as a fraction of the naphtha, to form benzene, and isvery endothermic, and formula (II) relates to the conversion of hexaneto form cyclohexane, which can be reformed in turn into benzene inaccordance with formula (I). The secondary reactions of formulae (III)and (IV), which may likewise take place during the preparation ofbenzene, are shown as well, and are especially exothermic reactions.

C₆H₁₂

C₆H₆+3.H₂  (I)

C₆H₁₄

C₆H₁₂+H₂  (II)

C₆H₁₂+2.H₂→2.C₃H₈  (III)

C₆H₁₄+H₂→2.C₃H₈  (IV)

The reaction according to formula (I), like that according to formula(II), is equilibrium-limited.

The benzene obtained from the reaction according to formula (I) forms akey starting product for further reaction to give, for example, theabovementioned products.

The controlled supply of heat in processes for obtaining benzene isimportant, since the position of the equilibrium in the aforementionedreaction according to formula (I) is heavily dependent on thetemperature of the reaction zone, and it is therefore possible tocontrol the yields and/or selectivities with respect to benzene by thismeans. In particular it is therefore possible at least partly tosuppress the unwanted secondary reactions of formulae (III-IV).

An uncontrolled drop in temperature as a result of the endothermicreaction according to formula (I) may therefore promote the formation ofmore or less large quantities of cyclohexane (according to formula I byback-reaction) and/or propane (according to formula III) and/or hexane(according to formula II by back-reaction), which is a disadvantage forthe subsequent use of the benzene, since such secondary constituentsmust first be separated off.

It is therefore advantageous to keep the temperature of the reactionzones in the course of the process controlled at a level which allowsrapid conversion with a minimization of the secondary reactions.

Accordingly, EP 0 601 398 A1 discloses a key influence on yield andconversion to target product exerted in the preparation of BTX aromatics(benzene, toluene and xylene) by the temperature level and by thecatalyst employed. In accordance with the disclosure content of EP 0 601398 A1, the temperature level at which the reaction is to be performedis essentially determined by the nature and composition of the naphthaused, which is typically characterized by its boiling point. Thisunderlines the importance of precise temperature control in suchprocesses.

EP 0 601 398 A1 also discloses how it is now customary to performcatalytic reformation processes in a plurality of reactors, connected inseries, containing catalysts in the form of a fixed bed.

EP 0 601 398 A1 discloses an isothermal procedure using a salt bath bymeans of which the temperatures of approximately 500° C. that aredisclosed in the process are brought about in the reaction zone. Anadiabatic regime is not disclosed. The catalyst used in the processdisclosed in EP 0 601 398 A1 is composed of a support material, which ispreferably alumina, on which there is located a layer of a platinumgroup metal with a promoter metal from group WB of the Periodic Table ofthe Elements.

The process as disclosed in EP 0 601 398 A1 is disadvantageous onaccount of the fact that the isothermal procedure disclosed is extremelycomplicated and therefore very expensive. In the production ofindustrial chemicals in particular, which includes the production ofbenzene, however, even slight process disadvantages have severeconsequences for the economics of the process as a whole, as is alsodisclosed by EP 0 601 398 A1.

The possibility of an adiabatic regime is disclosed by J.Ancheyta-Juarez et al. in “Modeling and simulation of four catalyticreactors in series for naphtha reforming” in Energy & Fuels (2001) 15:887-893.

Thus J. Ancheyta-Juarez et al. disclose how it can be advantageous toperform the reaction of naphtha to form (among other products) benzenein three to four, especially four, serial reaction zones, with thepossibility of intermediate cooling between the aforementioned reactionzones.

The yields of benzene (A₆) which can be achieved by means of the processpresented in the disclosure by J. Ancheyta-Juarez et al. are very low,with a fraction of only about 4 mol % as a proportion of the reactionproduct, thus making the process disadvantageous.

EP 1 251 951 (B1) discloses an apparatus and the possibility ofconducting chemical reactions in the apparatus, the apparatus beingcharacterized by a cascade of mutually contacting reaction zones andheat exchanger devices which are integrated materially with one another.The process to be conducted therein is characterized, therefore, by thecontact of the various reaction zones with a respective heat exchangerdevice, in the form of a cascade. A disclosure as to the possibility ofusing the device and the process for preparing benzene is absent.

It remains unclear, then, as to how, on the basis of the disclosurecontent of EP 1 251 951 (B1), a reaction of this kind is to be carriedout by means of the apparatus and of the process performed therein. Inparticular there is no disclosure of a process comprising endothermicreactions.

Furthermore, for reasons of consistency, it must be assumed that theprocess disclosed in EP 1 251 951 (B1) is performed in an apparatuswhich is identical or similar to the disclosure relating to theapparatus. The result of this is that, as a consequence of the extensivecontact between the heat exchange zones and the reaction zones, as perthe disclosure, a significant amount of heat takes place as a result ofthermal conduction between the reaction zones and the adjacent heatexchange zones.

The disclosure concerning the oscillating temperature profile can onlybe understood, then, to mean that the temperature peaks found here wouldbe sharper in the absence of this contact. A further indicator of thisis the exponential increase in the disclosed temperature profilesbetween the individual temperature peaks. These indicate that there is acertain heat sink present in each reaction zone, with a marked butlimited capacity, which is able to reduce the temperature increase insaid zone. It is never possible to rule out a certain dissipation ofheat (by radiation, for example); however, in the case of a reduction inthe possible dissipation of heat from the reaction zone, a linear ordegressive temperature profile would be suggested, since there is nosubsequent metering of reactants and hence, after their consumption byexothermic reaction, the reaction would become increasingly slow andhence the heat produced would go down.

EP 1 251 951 (B1) thus discloses multi-stage processes in cascades ofreaction zones from which heat is taken off in an undefined quantity bymeans of thermal conduction. The process disclosed, therefore, is notadiabatic and is disadvantageous insofar as precise temperature controlof the reaction is impossible. This applies especially to theundisclosed possibility of an endothermic reaction in the reactionzones.

On the basis of the prior art, therefore, it would be advantageous toprovide a process for preparing benzene that can be carried out insimple reaction apparatus and that enables precise, simple temperaturecontrol of the endothermic process, thereby allowing high conversions inconjunction with very high purities of the product, while meetingdesired yields and/or selectivities. Said simple reaction apparatuswould be readily transposable to the industrial scale, and inexpensiveand robust in all sizes.

As has just been shown, there have to date been no suitable processesapparent that allow this for the endothermic, catalytic gas phasereaction of naphtha to form benzene.

The object is therefore that of providing a process for endothermic,catalytic gas phase reaction of naphtha to benzene which can be carriedout with precise temperature control in simple reaction apparatus andwhich as a result allows high conversions in conjunction with highproduct purities.

It has surprisingly been found that a process for preparing benzene fromnaphtha in the presence of hydrogen in an endothermic, heterogeneouslycatalytic gas phase reaction, characterized in that it comprises 5 to 12serial reaction zones with adiabatic conditions, is able to achieve thisobject.

In connection with the present invention, benzene is a process gassubstantially comprising benzene. The benzene may also comprisefractions of hydrogen and further hydrocarbons.

In connection with the present invention, further hydrocarbons arecompounds present in the form of process gas composed of carbon,hydrogen and possibly oxygen. Essentially, however, such hydrocarbonsare composed of carbon and hydrogen. Such hydrocarbons are typicallyeither those which are introduced into the process of the invention asfurther constituents of the naphtha, or those which are formed as aresult of secondary reactions in the course of the process of theinvention, as for instance by the reactions according to formulae (IIIand IV).

Non-exhaustive examples of hydrocarbons which are introduced into theprocess of the invention as further constituents of the naphtha arenaphthalene, isopentane and toluene, for instance.

Non-exhaustive examples of hydrocarbons which are formed in the courseof the process of the invention by secondary reactions, as for instanceby the reactions according to formulae (III and IV), are hexane,cyclohexane and propane, for instance.

Naphtha identifies a mixture of hydrocarbons in the form of a processgas, as is general knowledge to the person skilled in the art. Inconnection with the process of the invention, naphtha is preferably amixture of hydrocarbons substantially comprising cyclohexane.

In connection with the present invention, hydrogen is a process gaswhich substantially comprises hydrogen. This hydrogen may be formed, forinstance, by the reactions according to formulae (I and II), or else maybe supplied as process gas to the process.

The supplying of hydrogen as a process gas into the process of theinvention is preferred. With particular preference, preheated hydrogenis supplied as process gas to the process.

Such supplying of hydrogen in particular is advantageous in that itallows the hydrogen to be used as a heat transfer medium in the process,for controlling the temperature. Furthermore, the hydrogen preventsdeposits of carbon products on the catalyst surfaces of the catalystslocated in the reaction zones (coking).

The identification “substantially” refers, in connection with thepresent invention, to a mass fraction and/or a molar fraction of atleast 80%.

The naphtha used in the process of the invention, its constituents, thehydrogen, the benzene and also the products of the process of theinvention are also referred to below collectively as process gases.

It follows from this that the entire process of the invention isperformed in the gas phase. If substances used in the process, such asthe hydrocarbons, for instance, are not in gaseous form at roomtemperature (23° C.) and ambient pressure (1013 hPa), it can be assumedbelow that, before or during their use in the process of the invention,such substances will be converted into the gas phase by an increase intemperature and/or reduction in pressure.

Besides the substantial components of the process gases, they may alsocomprise secondary components. Non-exhaustive examples of secondarycomponents which may be present in the process gases are argon, nitrogenand/or carbon dioxide, for instance.

In accordance with the invention the implementation of the process underadiabatic conditions means that substantially neither heat is activelysupplied nor heat withdrawn from the outside to/from the reaction zone.It is common knowledge that complete insulation from ingress or egressof heat is possible only by complete evacuation, with the possibility ofheat transfer by radiation being ruled out. In connection with thepresent invention, therefore, adiabatic means that no measures are takento supply or remove heat.

In one alternative embodiment of the process of the invention, however,heat transfer may be reduced, for example, by insulation usingconventional insulating means, such as polystyrene insulants, forexample, or else by sufficiently large distances from heat sinks or heatsources, the insulation means being air.

An advantage of the adiabatic regime of 5 to 12 serial reaction zones inaccordance with the invention as compared with a non-adiabatic regime isthat in the reaction zones there is no need to provide means for heatremoval, a fact which results in a considerable simplification inconstruction. As a result, in particular, there are simplifications inthe manufacturing of the reactor and also in the scaleability of theprocess, and there is an increase in the reaction conversions.

A further advantage of the process of the invention is the possibilityof very precise temperature control, as a result of the narrowstaggering of adiabatic reaction zones. By this means it is possible ineach reaction zone to set and control a temperature which isadvantageous in the progress of the reaction.

The catalysts used in the process of the invention are typicallycatalysts composed of a material which as well as its catalytic activityfor the reaction according to formula (I) is characterized by sufficientchemical resistance under the conditions of the process and also by ahigh specific surface area.

Catalyst materials which are characterized by such chemical resistanceunder the conditions of the process are, for example, catalystscomprising platinum and/or rhenium.

Preferred catalyst materials are composed of equal weight fractions ofrhenium and platinum.

These catalysts may be applied on support materials. Such supportmaterials typically comprise alumina and/or titanium dioxide. Preferenceis given to alumina support materials.

Particularly preferred catalysts are composed of rhenium and platinumapplied at the same weight fraction on an alumina support. Methods ofproducing such catalysts are general knowledge to a person skilled inthe art, from EP 0 601 398 A1, for instance.

Specific surface area in connection with the present inventionidentifies the surface area of the catalyst material which can bereached by the process gas, based on the mass of catalyst materialemployed.

A high specific surface area is a specific surface area of at least 1m²/g, preferably of at least 10 m²/g.

The catalysts of the invention are located in the reaction zones in eachcase and may be present in all conventional presentation forms, e.g.fixed bed, moving bed.

The presentation form is preferably that of a fixed bed.

The fixed bed arrangement comprises a catalyst bed in the actual sense,i.e. loose, supported or unsupported catalyst in any desired form, andalso in the form of suitable packings. The term catalyst bed as usedherein also encompasses coherent regions of suitable packings on asupport material or structured catalyst support. Examples of such wouldinclude ceramic honeycomb supports for coating, having comparativelyhigh geometric surface areas, or corrugated layers of metal wire meshwith catalyst granules, for example, immobilized thereon. In connectionwith the present invention, the presence of the catalyst in monolithicform is viewed as a special form of packing.

Where a fixed bed arrangement of the catalyst is used, the catalyst ispreferably in beds of particles having average particles sizes of 1 to10 mm, preferably 2 to 8 mm, more preferably of 3 to 7 mm.

Likewise with preference the catalyst in the case of a fixed bedarrangement is in monolithic form. In the case of a fixed bedarrangement particular preference is given to a monolithic catalystwhich comprises the aforementioned metals, rhenium and platinum, inequal weight fractions on an alumina support.

Likewise particularly preferred is a fixed bed arrangement havingparticle beds, having average particles sizes of 1 to 10 mm, preferably2 to 8 mm, more preferably of 3 to 7 mm, the particles being aluminaparticles to which the aforementioned metals, rhenium and platinum, havebeen applied in equal weight fractions.

If a catalyst is used in monolithic form in the reaction zones, then, ina preferred development of the invention, the catalyst present inmonolithic form is provided with channels through which the processgases flow. Typically the channels have a diameter of 0.1 to 3 mm,preferably a diameter of 0.2 to 2 mm, more preferably of 0.5 to 1.5 mm.

Where a moving bed arrangement of the catalyst is used, the catalystpreferably takes the form of loose beds of particles, of the kindalready described in connection with the fixed bed arrangement.

Beds of such particles are advantageous because the size of theparticles have a high specific surface area of the catalyst material inrelation to the process gases and it is therefore possible to achieve ahigh conversion rate. Accordingly, the limitation of mass transport inthe reaction by diffusion can be minimized. At the same time, however,the particles are also not so small that increased pressure drops occurdisproportionately when the gases flow through the fixed bed. The rangesof particle sizes specified in the preferred embodiment of the process,comprising a reaction in a fixed bed, are therefore an optimum betweenthe achievable conversion from the reactions according to formulae (Iand II) and the pressure drop generated when the process is implemented.Pressure drop is coupled directly with the necessary energy, in the formof compressor output, and so a superproportional increase in the latterwould result in an uneconomic process regime.

In one preferred embodiment of the process of the invention theconversion takes place in 6 to 10, more preferably 6 to 8, serialreaction zones.

A preferred further embodiment of the process is characterized in thatthe process gas emerging from at least one reaction zone is subsequentlypassed through at least one heat exchange zone downstream of saidreaction zone.

In one particularly preferred further embodiment of the process eachreaction zone is followed by at least one, preferably exactly one, heatexchange zone through which the process gas emerging from the reactionzone is passed.

These reaction zones may be disposed either in one reactor or, individed form, in two or more reactors. The arrangement of the reactionzones in one reactor leads to a reduction in the number of apparatusesused.

The individual reaction zones and heat exchange zones may also bearranged in one reactor or, in divided form, in any desired combinationsof reaction zones with heat exchange zones in two or more reactors.

Where reaction zones and heat exchange zones are present in one reactor,then, in an alternative embodiment of the invention, there is a heatinsulation zone located between these zones in order to allow theadiabatic operation of the reaction zone to be maintained.

In addition it is possible for certain of the serial reaction zones tobe replaced or supplemented, independently of one another, by one ormore parallel reaction zones. The use of parallel reaction zones makesit possible in particular to exchange or supplement the zones duringongoing, continuous operation of the process overall.

Parallel and serial reaction zones may in particular also be combinedwith one another. With particular reference, however, the process of theinvention features exclusively serial reaction zones.

The reactors used preferably in the process of the invention may becomposed of simple vessels with one or more reaction zones, of the kinddescribed, for example, in Ullmann's Encyclopedia of IndustrialChemistry (Fifth, Completely Revised Edition, Vol. B4, pages 95-104,pages 210-216), it being possible for heat insulation zones to beprovided additionally between each of the individual reaction zonesand/or heat exchange zones.

In one alternative embodiment of the process, therefore, there is atleast one heat insulation zone located between a reaction zone and aheat exchange zone. Preferably there is a heat insulation zone locatedaround each reaction zone.

The catalysts or fixed beds thereof are mounted in a conventional way onor between gas-permeable walls encompassing the reaction zone of thereactor. In the case of thin fixed beds in particular, technical devicesfor uniform gas distribution may be fitted upstream of the catalystbeds. These devices may be perforated plates, bubble trays, valve traysor other internals which, by generating a low but uniform pressure drop,produce uniform entry of the process gas into the fixed bed.

In one preferred embodiment of the process the entry temperature of theprocess gas entering one reaction zone is from 740 to 790 K, preferablyfrom 750 to 780 K, more preferably from 755 to 775 K.

In a further preferred embodiment of the process the absolute pressureat the entry of the first reaction zone is between 10 and 40 bar,preferably between 15 and 35 bar, more preferably between 20 and 30 bar.

In yet another preferred embodiment of the process the residence time ofthe process gas in all the reaction zones together is between 0.5 and 30s, preferably between 1 and 20 s, more preferably between 5 and 15 s.

The naphtha and, where appropriate, the hydrogen are preferably suppliedonly ahead of the first reaction zone. This has the advantage that thewhole of the process gas is available for the accommodation of heat ofreaction in all the reaction zones. Such a procedure additionallyenables the space-time yield to be increased, or the mass of catalystrequired to be reduced. It is, however, also possible to meter naphthaand, where appropriate, hydrogen into the process gas ahead of one ormore of the reaction zones that follow the first reaction zone, ifneeded. The supply of these process gases between the reaction zone isan additional way of controlling the temperature of the conversion, ifthey are preheated.

In preferred embodiments of the process of the invention the molar ratioof hydrogen to hydrocarbons present in the naphtha is set in ranges from3 to 9, preferably from 4 to 8, more preferably from 5 to 7 mol ofhydrogen per mole of hydrocarbon in the naphtha.

The advantages of such supply of hydrogen have already been elucidated.They apply in particular in connection with the supply of an excess.

The person skilled in the art is aware of suitable means for determiningthe molar amounts of hydrocarbons in a process gas, such as naphtha. Onenon-exhaustive example is quantitative analysis by means of gaschromatography. If the molar composition of the naphtha process gas isknown, the molar ratio of hydrogen to it can be set by simple setting ofthe volume flow ratio of the naphtha and hydrogen process gases.

In a further preferred embodiment of the process of the invention theprocess gas is heated after at least one of the reaction zones used,more preferably after each reaction zone. This is done by passing theprocess gas, following exit from a reaction zone, through one or more ofthe abovementioned heat exchange zones which are located downstream ofthe respective reaction zones. These zones may be configured as heatexchange zones in the form of heat exchangers known to the personskilled in the art, such as shell-and-tube, plate, annular-groove,spiral, ribbed-tube or micro-type heat exchangers, for example. The heatexchangers are preferably microstructured heat exchangers.

Microstructured in connection with the present invention means that theheat exchanger, for the purpose of heat transfer, comprisesfluid-carrying channels which are characterized in that they have ahydraulic diameter of between 50 μm and 5 mm. The hydraulic diameter iscalculated from four times the flow-traversed cross-sectional area ofthe fluid-carrying channel, divided by the circumference of the channel.

In one particular embodiment of the process the process gas is heated inthe heat exchange zones by condensation of a heat transfer medium.

Within this particular embodiment it is preferred to performcondensation, preferably partial condensation, on the side of theheating medium in the heat exchangers which constitute the heat exchangezones.

Partial condensation in connection with the present invention means acondensation in which the heating medium used is a substance in the formof a gas/liquid mixture, and in which, following heat transfer in theheat exchanger, this substance is still in the form of a gas/liquidmixture.

Performing a condensation is particularly advantageous since it meansthat the coefficient of heat transfer to the process gases from theheating medium that can be achieved becomes particularly high and hencethat it is possible to achieve efficient heating.

Performing a partial condensation is particularly advantageous becauseit means that the delivery of heat by the heating medium no longerresults in a temperature change to the heating medium, but insteadmerely shifts the gas/liquid equilibrium. As a result of this, theprocess gas is heated towards a constant temperature over the entireheat exchange zone. This in turn reliably prevents the incidence ofradial temperature profiles in the flow of the process gases, therebyimproving the control of the reaction temperatures in the reaction zonesand, in particular, preventing the development of instances of localoverheating as a result of radial temperature profiles.

In an alternative embodiment, instead of a condensation/partialcondensation, it is also possible to provide a mixing zone before theentry of a reaction zone, in order to unify any radial temperatureprofiles formed in the course of heating in the flow of the processgases by mixing transverse to the main flow direction.

In one preferred embodiment of the process the succession of reactionzones are operated with an average temperature rising or falling fromreaction zone to reaction zone. This means that, within a sequence ofreaction zones, the temperature may be made both to rise and to fallfrom reaction zone to reaction zone. This can be brought about, forexample, by means of control of the heat exchange zones inserted betweenthe reaction zone. Further possibilities for setting the averagetemperature are described below.

The thickness of the flow-traversed reaction zones may be the same ordifferent and is a function of laws which are general knowledge to theperson skilled in the art and relate to the above-described residencetime and the volumes of process gas processed in each case. The massflows of process gas that can be processed in accordance with theinvention, relative to the mass of catalyst used (and also called WHSV,Weight-Hourly Space Velocity), is typically between 28 and 42 h⁻¹,preferably between 30 and 40 h⁻¹, more preferably between 33 and 38 h⁻¹.

The maximum exit temperature of the process gas from the first reactionzone is typically in the region of the entry temperature, since thereactions according to formulae (III) and (W) are exothermic reactions.Particularly in the case of exit from the final reactions, in which alarge amount of benzene has already been formed and therefore theespecially endothermic reaction according to formula (I) is losing itsinfluence, these temperatures may also be situated within a range from770 to 820 K, preferably from 775 to 795 K, more preferably from 780 to785 K.

The person skilled in the art is able freely to determine the entrytemperature of the subsequent reaction zones by means of the measuresbelow in accordance with the process of the invention.

Temperature control in the reaction zones is accomplished preferably byat least one of the following measures: sizing of the adiabatic reactionzone, control of the heat supply between the reaction zones, addition offurther process gas between the reaction zones, molar ratio ofreactants/excess of hydrogen used, addition of secondary constituents,especially nitrogen, carbon dioxide, ahead of and/or between thereaction zones.

The composition of the catalysts in the reaction zones of the inventionmay be the same or different. In one preferred embodiment the samecatalysts are used in each reaction zone. An advantageous alternative isto use different catalysts in the individual reaction zones.

Thus it is possible in particular in the first reaction zone, when theconcentration of the reactants is still high, to use a less activecatalyst, and to increase the activity of the catalyst from reactionzone to reaction zone in the further reaction zones. The catalystactivity can also be controlled by dilution with inert materials and/orsupport material.

With the process of the invention it is possible to prepare 1 kg/h to 50kg/h, preferably 5 kg/h to 30 kg/h, more preferably 10 kg/h to 20 kg/hof benzene per kg of catalyst.

The process of the invention is therefore distinguished by highspace-time yields, in conjunction with a reduction in apparatus sizesand also with a simplification of the apparatus and/or reactors. Thissurprisingly high space-time yield is made possible through theinteraction of the inventive and preferred embodiments of the newprocess. The interaction of staggered adiabatic reaction zones withinterposed heat exchange zones and the defined residence times, inparticular, allows precise control of the process and the resulting highspace-time yields, and also a reduction in the by-products formed, suchas carbon dioxide, for instance.

The present invention is illustrated with reference to the figures,though without being restricted thereto.

FIG. 1 shows reactor temperature (T) and molar mass flows of benzene (U)over a length (L) of 11 m of reaction zones each with downstream heatexchange zones (in accordance with Example 1), the lengths of the heatexchange zones being assumed ideally to be zero, since no conversion isto take place here.

The present invention is further illustrated by the following example,but without being limited thereto.

EXAMPLES

Gaseous naphtha and hydrogen are supplied to the process as processgases in a molar ratio of 7.77. The process is operated in a total ofsix fixed catalyst beds of rhenium and platinum, each at 0.29% byweight, on an alumina support, in other words in six reaction zones.

After each reaction zone there is a heat exchange zone in which theexiting process gas is heated again before entering the next reactionzone.

The absolute entry pressure of the process gas directly ahead of thefirst reaction zone is 25 bar. The length of the fixed catalyst beds,and therefore of the reaction zones, varies from reaction zone toreaction zone, beginning from 0.15 m in the first reaction zone throughto 6 m in the sixth reaction zone. The precise links of the reactionzone are summarized in Table 1. The activity of the catalyst used isunvarying over the reaction zones. No process gas is metered in ahead ofthe individual reaction zones. The WHSV is 35 h⁻¹.

TABLE 1 Lengths of the reaction zones Reaction Length zone [#] [m] 10.15 2 0.35 3 1 4 1.5 5 2 6 6 Σ 11.0

The results are shown in FIG. 1. Varying the cumulative length of thereaction zones is plotted on the x-axis, so that it is possible to see aspatial course of the developments in the process; the heat exchangezones are disregarded. On the left-hand, y-axis, the temperature of theprocess gas is indicated. The temperature profile over the individualreaction zones is depicted as a thick, continuous line. As a result ofthe idealized assumption of the length of the heat exchange zones as 0m, there are discontinuities in the temperature profile. On theright-hand y-axis the cumulative molar flow of benzene in the processgas over the reaction path is indicated. Its profile over said path isdepicted as a thin continuous line.

It can be seen that the entry temperature of the process gas ahead ofthe first reaction zone is approximately 775 K. As a result of thesubstantially endothermic reaction to form benzene under adiabaticconditions the temperature in the first reaction zone drops to about 760K, before in the downstream heat exchange zone the process gas isreheated to the aforementioned 775 K. As a result of endothermicadiabatic reaction, the temperature in the second reaction zone drops toabout 750 K. The sequence of cooling as a result of endothermic,adiabatic reaction, and heating continues, with changes in exittemperatures, after the respective reaction zones, the entry temperaturebeing re-established at the desired 775 K in each of the heat exchangezones.

A conversion of cyclohexane and hexane of approximately 60% is obtained.The space-time yield achieved, based on the mass of catalyst employed,is approximately 15 kg_(benzene)/kg_(cat)h.

1. Process for preparing benzene from naphtha in the presence ofhydrogen in an endothermic, heterogeneously catalytic gas phasereaction, which comprises 5 to 12 serial reaction zones with adiabaticconditions.
 2. Process according to claim 1, wherein the conversiontakes place in 6 to 10 serial reaction zones.
 3. Process according toclaim 1, wherein the entry temperature of the process gas entering thefirst reaction zone is 740 to 790 K.
 4. Process according to claim 1,wherein the absolute pressure at the entry of the first reaction zone isbetween 10 and 40 bar.
 5. Process according to claim 1, wherein theresidence time of the process gas in all reaction zones is between 0.5and 30 s.
 6. Process according to claim 1, wherein the catalysts arepresent in a fixed bed arrangement.
 7. Process according to claim 6,wherein the catalysts are present in the form of monoliths.
 8. Processaccording to claim 7, wherein the monolith comprises channels having adiameter of 0.1 to 3 mm.
 9. Process according to claim 1, wherein thecatalysts are present in beds of particles having average particlessizes of 1 to 10 mm.
 10. Process according to claim 1, wherein at leastone reaction zone is followed by at least one heat exchange zone throughwhich the process gas is passed.
 11. Process according to claim 10,wherein each reaction zone is followed by at least one heat exchangezone through which the process gas is passed.
 12. Process according toclaim 1, wherein between a reaction zone and a heat exchange zone thereis at least one heat insulation zone.
 13. Process according to claim 12,wherein around each reaction zone there is a heat insulation zone.